Selective Conversion of Paraffinic Naphtha to Propylene in the Presence of Hydrogen

ABSTRACT

The invention relates to a process of catalytic conversion by hydrocracking of paraffinic and naphthenic hydrocarbons from a naphtha feedstock ( 1 ) to propylene, the process comprising the steps of providing a naphtha feedstock ( 1 ) containing one or more paraffins comprising 4 to 10 carbon atoms; and contacting said naphtha feedstock ( 1 ) with a catalyst composition in the presence of hydrogen in a reaction zone under hydrocracking conditions; wherein the catalyst composition consists of one or more zeolite catalysts comprising acid 10-membered ring channels.

FIELD OF THE INVENTION

The invention relates to a process for producing propane from a naphthafeedstock comprising paraffinic and naphthenic hydrocarbons, and tofurther produce propylene from the propane recovered. The invention alsorelates to the use of catalyst compositions to improve propaneselectivity in a process for producing propane from a naphtha feedstockcomprising paraffinic and naphthenic hydrocarbons.

BACKGROUND OF THE INVENTION

World demand for propylene is expected to continuously grow up at anaverage annual rate of 4-5%. Currently, propylene is mainly produced asa by-product of liquefied petroleum gas (LPG)/naphtha steam crackers andFluid Catalytic Cracking (FCC) units. With the startup of additionalethane-based ethylene capacities (lighter fractions like ethane or LPGare considered very advantageous naphtha feedstock), the production ofpropylene and aromatics via the steam cracker declines. Furthermore, thedemand for propylene is actually growing faster than that of ethylene.To some extent, the propylene production can be optimized bytransforming an FCC unit into a petrochemical FCC after a major revamp.However, a significant amount of dry gases and low value of by-productswill be produced. In addition, an important revamp of the existing FCCunit and significantly higher process severity would be required tocrack the naphtha to propylene and ethylene.

All the aforementioned factors have created an imbalance of supply anddemand for propylene: a gap is being established between the availablepropylene supplies to meet ongoing demand growth. While the markets haveevolved to the point where modes of propylene by-product production canno longer satisfy the demand for propylene, the traditionalclassification of propylene as a “by-product” begun to evolve into moreof classification as a “coproduct” or even a “primary product”.

Propane dehydrogenation (PDH) is the fastest on-purpose growingtechnology today to bridge the gap. The technology offers a mean ofproducing propylene as a single product near consumer markets. PDHsystem is suitable as stand-alone facilities not requiring integrationwith a nearby olefin or refinery unit but requires a secured supply of abig amount of propane (>0.5 Mt/y).

The PDH units used to be only appropriate for a limited amount ofgeographical regions, where propane was highly available. Today, themajority of the PDH projects are based on the imported propane from theUnited States of America (US). A direct propane import from theproducers in the US or the Middle East has many limits on volume andrequires sophisticated logistics. On the contrary, the market expects animportant surplus in the availability of naphtha due to a massive shiftto ethane as feedstock to steam-crakers and low oil prices. Therefore,there is a need for the development of a technology to transform naphthato propylene with significant yield advantages versus naphtha steamcracker and petro FCC.

The majority of the catalytic processes to crack paraffinic naphtha isdedicated to direct production of propylene in absence of hydrogen at atemperature higher than 550° C. and low partial pressure of hydrocarbonbelow atmospheric. These conditions reduce hydrogen transfer reactionsand favor propylene and aromatics vs propane (D5943, Applied CatalysisA: General 398 (2011) 1-17). Propane is formed but as an undesirableside product. One should mention also that under high-temperatureconditions, a lot of ethylene is produced which, in the context ofgrowing ethylene supply from ethane crackers, reduces significantlyattractiveness of the process.

The isomerization and hydrotreating processes of naphtha are well knownto produce clean fuel from linear paraffinic species existing in lightnaphtha. Both processes occur in the presence of hydrogen at moderatepressure on the metal-containing catalysts with a purpose to increasefuel qualities. Light formation is big drawback for both reactions. Apost-reforming process known as selectoforming was commercialized in the1960s for raising the octane rating of reformates while producingpropane as a by-product (Chen et al., 1968; Burd and Maziuk, 1972).

Some articles report naphtha hydrocracking, for example: “High Yields ofLPG Via n-Hexane Hydrocracking Using Unloaded Acidic Zeolite Catalysts”in Journal Petroleum Science and Technology, Volume 33, 2015—Issue 12,includes investigating the hydrocracking of n-hexane, as a low-octanenaphtha component to high-octane gaseous motor fuel (LPG) in a pulseflow atmospheric microreactor using untreated and steam-treated H-MOR,H-BEA, or H-ZSM-5 zeolite catalysts. All zeolites catalysts weremetal-free and their bi-functionality depended only on the Brønstedzeolitic acid sites. Zeolites catalysts H-ZSM-5 and St-H-ZSM-5 acquirevery low catalytic activities mainly due to their narrow pore structure,as well as due to their partial pore filling by Al debris in case ofSt-H-ZSM-5.

In “Conversion of light naphthas over sulfided nickel erionite” inIndustrial and Engineering Chemistry Research Volume 32, Issue 6, 1993,Pages 1003-1006, a natural zeolite erionite has been exchanged withammonium and nickel salts to yield a Ni/H erionite catalyst that isactive and stable for selectively hydrocracking only the n-paraffinsfrom light straight-run naphthas. The primary product is a C5+ liquidthat has 15-20 octane numbers higher than the feed and a propane- andbutane-rich gas by-product. Results from a 110-day pilot plant rundemonstrated that a catalyst life of more than 1 year should bepossible. Naphthenes, aromatics, and isoparaffins are neither producednor consumed in this process, resulting in a C5+ liquid product that islower in benzene and total aromatics than attainable by catalyticreforming of these feeds. Although no further work is planned with thiscatalyst, a naphtha-upgrading process based on shape-selective zeolitichydrocracking could provide an attractive alternative to catalyticreforming or isomerization for these hard-to-upgrade naphthas. It shouldbe particularly attractive in areas where the propane and butaneby-products have good value.

WO2012/071137 describes the use of a catalyst comprising at least onezeolite having 10-membered ring channels and at least one group VIb,VIIb and/or VIII metal, in the presence of hydrogen at elevatedtemperature and elevated pressure. The process allowed to convert atleast 40 wt % of the paraffins comprising from 4 to 12 carbon atomsbased on the total weight of paraffins in the feed to ethane and/orpropane to obtain a hydrocracked gas cracker naphtha feedstockcomprising ethane and/or propane. The process was focussed on maximizingthe production of ethane. The drawback of the invention is an importantformation of methane due to the presence of the noble metals in thecatalyst at about 6-20 wt %.

US 2017/058210 discloses a process for producing BTX comprisingpyrolysis, aromatic ring opening and BTX recovery.

US 2016/369191 discloses a process for cracking a hydrocarbon feedstockin a steam cracker unit where a liquid hydrocarbon feedstock is fed to ahydrocracking unit, the stream hydrocracked are then separated into ahigh content aromatics stream and a gaseous stream comprising C2-C4paraffins, hydrogen and methane, the C2-C4 paraffins are then separatedform said gaseous stream and fed to a steam cracker unit.

US 2017/369795 discloses a process for producing C2 and C3 hydrocarbons,comprising a) subjecting a mixed hydrocarbon stream to firsthydrocracking in the presence of a first hydrocracking catalyst toproduce a first hydrocracking product stream; and b) subjecting thefirst hydrocarbon product stream to C4 hydrocracking.

WO 98/56740 discloses a process for improving the conversion of ahydrocarbon feedstock to light olefins comprising the steps of thermallyconverting the hydrocarbon feedstock to produce an effluent; quenchingthe effluent to produce a quenched effluent; and contacting the quenchedeffluent with a light olefin-producing cracking catalyst.

It was discovered that the naphtha may selectively be transformed intopropane on the metal free catalyst (i.e. noble and transition metal freecatalyst) with very low selectivity to methane (from 0.3 to 1.5 wt %CH4). This opens an opportunity for a high efficient process ofproduction of propylene from naphtha via a combination of hydrocrackingof naphtha to propane followed by dehydrogenation of propane topropylene. The dehydrogenation process produces hydrogen, which could beused at least partially for the first step.

SUMMARY OF THE INVENTION

According to a first aspect, the invention provides a process ofcatalytic conversion by hydrocracking of a paraffinic and naphthenichydrocarbons from a naphtha feedstock to propylene, wherein the processcomprises the following steps:

-   -   a) providing a naphtha feedstock containing one or more        paraffins comprising 4 to 10 carbon atoms with preferably no        olefins i.e. an olefin content of less than 1 wt % preferably        less than 0.1 wt % even more preferably less than 0.01 wt %        relating to the total weight of said naphtha feedstock; and    -   b) submitting said naphtha feedstock to a hydrocracking step by        contacting said naphtha feedstock with a catalyst composition in        the presence of hydrogen in a reaction zone under hydrocracking        conditions to produce an effluent;    -   c) submitting the effluent to a separation step to recover        propane; and    -   d) submitting said propane to a step of dehydrogenation into        propylene in a propane dehydrogenation reactor;    -   e) collecting hydrogen (21) produced in the step of        dehydrogenation into propylene (23), and recycling said        hydrogen (21) back to the hydrocracking reaction zone in order        to perform the hydrocracking step (3)        and wherein the catalyst composition of the hydrocracking step        comprises one or more zeolite catalysts comprising an acid        10-membered ring channels.

Surprisingly, it was found by the inventors that the catalystcomposition comprising acid 10-membered ring channels was particularlyresistant to deactivation. Indeed the hydrogen recycled from thedehydrogenation step to the hydrocracking reaction zone is notnecessarily of the highest purity. This hydrogen can contain methane orother impurities like sulphur at relatively high concentration thatdeactivate traditional hydrocracking catalyst. The use of the catalystcomposition comprising acid 10-membered ring channels is resistant todeactivation induced by the methane contained in hydrogen. Such catalystallows avoiding expensive purification of the hydrogen. The hydrogenfeed stream can for instance contain up to 5 wt % of methane, preferablyup to 10 wt %, even more preferably up to 20 wt % of methane based onthe total weight of the hydrogen feed stream. Other impurities may alsobe present such as sulphur compounds, such as H₂S, at concentration forinstance up to 0.1% wt, preferably 1% wt, even more preferably up to 5wt % based on the total weight of the hydrogen feed stream. In apreferred embodiment, the one or more acid zeolite catalysts or catalystcomposition are metal-free, containing less than 1000 ppm of noble metaland less than 1% of transition metals. The content of the noble metalsis below 1000 ppm, preferably below 500 ppm, more preferably at most 250ppm even more preferably at most 50 ppm wt, the most preferred being nonoble metal at all, that is to say below the detection limit. Inexceptional cases, traces of noble metals (Pt, Pd) may be introduced toincrease the stability of the zeolite and hydrogenate the cokeprecursors. However, higher concentration of the metals may lead tochanges of the reaction mechanism. Presence of a significant amount ofnoble metals (>0.2 wt %) on the catalyst leads to a different reactionmechanism, which involves activation of the hydrogen on the metal(hydrogen spillover) followed by subsequent cracking. This leads to adifferent selectivity pattern and formation of higher amount of ethaneand methane. It is therefore preferred to have no noble metals at all.

As used herein the terms “metal-free” indicate that the one or more acidzeolite catalysts are also devoid of any transition metal selected fromthe groups of Fe, Ni, Co, W, Mo, Ga, Zn. The content of the transitionmetals is below 1.0 wt %, preferably below 0.5 wt % more preferablybelow 0.05 wt % even more preferably below 0.01 wt %. Traces of thesemetals may be present on the catalyst as impurities from the binder,e.g. a component of the clays. Surprisingly, it was found by theinventors that metal-free zeolite catalysts comprising acid 10-memberedring channels showed a high selectivity to propane and at the same time,stable performance in the presence of the aromatics presents in thenaphtha feedstock. The use of said catalysts in catalyst compositionsallows a direct cracking in presence of hydrogen of a naphtha feedstockcomprising paraffinic and naphthenic hydrocarbons without deactivationof the catalyst.

It is noted that U.S. Pat. No. 4,061,690 discloses a method of catalyticconversion of a butane cut to propane by means of a catalyst consistingof acid mordenite in the presence of hydrogen, with a partial hydrogenpressure higher than 0.5 MPa. The results show that the use of a largepore mordenite exchanged by protons (Zeolons H) at 400° C. and at apressure of 3 MPa had a propane yield of 75.0 wt % for a rate ofconversion of 92%, which corresponds to a propane selectivity ofapproximately 69%. The weighted analysis of the catalyst was SI: 40.6%;Al 6.2%; Na: 0.2%. The feedstock comprised a mixture of n-butane (60 wt%) and of iso-butane (40 wt %). These results were promising, however,it was found that MOR catalysts, having 12-membered ring channels,deactivate in the presence heavier feedstock with the number of atomssufficient to form aromatics or even containing the aromatics (thenaphtha feedstock). Thus, it was not possible to use MOR catalysts indirect hydrocracking of naphtha (DCN). By way of examples, it was shownthat MOR deactivates very fast naphtha feedstock and is not suitable.

With preference, one or more of the following embodiments can be used tobetter define the inventive process:

-   -   The naphtha feedstock comprises C4-C10 hydrocarbons with a        potential presence of the aromatics. The feedstock contains        preferably no olefins i.e. an olefin content of less than 0.1 wt        % relating to the total weight of said naphtha feedstock. The        C5+ hydrocarbons lead to a significant deactivation of the        12-members ring zeolites catalysts. The small pore erionite        could crack only a part of the feedstock. The solutions allow        treating the total amount of the feedstock without significant        deactivation of catalyst.    -   The naphtha feedstock comprises at least 10 wt % of naphthenes        as based on the total weight of the naphtha feedstock.    -   The naphtha feedstock comprises at least 2.0 wt % of aromatics        of five or more carbon atoms as based on the total weight of the        naphtha feedstock, preferably at least 2.2 wt %.    -   The naphtha feedstock comprises at most 10.0 wt % of aromatics        of five or more carbon atoms as based on the total weight of the        naphtha feedstock, preferably at most 9.0 wt %.    -   The catalyst composition is metal-free, containing less than        1000 wt ppm, preferably less than 50 ppm wt, more preferably at        most 5 ppm wt of noble metal and less than 1 wt %, preferably        less than 0.05 wt % even more preferably below 0.01 wt % of        transition metals as based on the total weight of the catalyst        composition. The catalyst composition is not sensitive to the        presence of sulfur and provides low C1-02 selectivity.    -   The hydrocracking conditions of the hydrocracking step comprise        the naphtha feedstock being contacted with the catalyst        composition at a temperature ranging from 200° C. to 600° C.,        preferably ranging from 250° C. to 550° C., more preferably        ranging from 300° C. to 450° C.    -   The hydrocracking conditions of the hydrocracking step comprise        the naphtha feedstock being contacted with the catalyst at a        pressure ranging from 1 to 10 MPa, preferably in the range of 2        to 6 MPa, more preferably from 3 to 5 MPa.    -   The hydrocracking conditions of the hydrocracking step comprise        the naphtha feedstock being contacted with the catalyst at a        WHSV (feed) of at least 0.1 h⁻¹, preferably is ranging from 0.1        h⁻¹ to 10.0 h⁻¹, more preferably from 0.5 h⁻¹ to 8.0 h⁻¹, even        more preferably from 1.0 h⁻¹ to 6.0 h⁻¹, and most preferably        from 1.5 h⁻¹ to 5.0 h⁻¹.    -   In the hydrocracking step, hydrogen is provided to the naphtha        feedstock at a molar ratio H₂/Naphtha ranging from 1000:1 to        1:1, preferably from 100:1 to 1:1, more preferably from 20:1 to        1:1. The hydrocracking step is particularly advantaging in that        hydrogen with a low purity can preferably be used. That is to        say that the hydrogen feed stream can for instance contain up to        5 wt % of methane, preferably up to 10 wt %, even more        preferably up to 20 wt % of methane based on the total weight of        the hydrogen feed stream. Other impurities may also be present        such as sulphur compounds, such as H₂S, at concentration up to        0.1% wt, preferably 1% wt, even more preferably up to 5 wt %        based on the total weight of the hydrogen feed stream.    -   The hydrocracking step is performed in a single reactor.

With preference, one or more of the following embodiments can be used tobetter define the catalyst used in the inventive process:

-   -   The one or more zeolite catalysts have a Si/Al molar ratio        ranging from 10 to 100, preferably from 20 to 80, more        preferably from 30 to 60.    -   At least 50 wt % of said one or more zeolite catalysts        comprising an acid 10-membered ring channels are in their        hydrogen form as based on the total weight of the zeolites        catalysts.    -   The catalyst composition comprises one or more zeolites        catalysts comprising an acid 10-membered ring channels selected        from the list comprising ZSM-5, silicalite-1, ZSM-11,        silicalite-2, SSZ-46, MCM-68, CIT-1, SSZ-33, ZSM-8, Ferrierite,        FU-9, ZSM-35, ZSM-23, ZSM-22, Theta-1, NU-10, ZSM-50, EU-1,        ZSM-57, SAPO-11 and ZSM-48 being preferably under their acidic        form or H-form.    -   Optionally, the one or more zeolites catalysts comprising an        acid 10-membered ring channels may be modified by phosphorous,        steaming, leaching, mesoporization, dealumination or        modification with alkali-earth or rare-earth metals.    -   The preferred zeolite catalyst structure is MFI (H-ZSM-5).    -   The catalyst composition comprises a binder selected from        silica, alpha-alumina, gamma-alumina, clays, alumina phosphates,        mullite, zirconia, titania, yttria, silicon nitride, silicon        carbide, iron, bronze and stainless steel, glass, and carbon,        preferably the binder is alumina.

In an embodiment, the step of dehydrogenation into propylene isperformed at a temperature ranging from 500 to 800° C. and a partialpressure of propylene below one atmosphere and/or the process furthercomprises a step of collecting hydrogen produced in the step ofdehydrogenation into propylene, and a step of recycling said hydrogenback to the hydrocracking reaction zone in order to perform thehydrocracking step.

In an embodiment, the separation step comprises recovering a C4paraffins fraction from the effluent of step b) and a step of recyclingback the C4 paraffins fraction to the hydrocrakrting reaction zone,and/or the separation step comprises recovering a C5+ hydrocarbonfraction from the effluent of step b) and a step of valorization of saidC5+ hydrocarbon fraction as gasoline.

In an embodiment, the process comprises a step of recovering theunconverted propane after the step of dehydrogenation into propylene andrecycling it to the propane dehydrogenation reactor to be submitted to afurther step of dehydrogenation into propylene.

According to a second aspect, the invention provides the use of one ormore zeolite catalysts comprising acid 10-membered ring channels in aprocess as defined according to the first aspect of catalytic conversionby hydrocracking of paraffinic and naphthenic hydrocarbons from anaphtha feedstock to propane, and further dehydrogenation of saidpropane to propylene, wherein the one or more zeolites catalystscomprising an acid 10-membered ring channels are metal-free.

With preference, the one or more zeolites catalysts comprising an acid10-membered ring channels comprise or are zeolites catalysts of theMFI-type, preferably the one or more zeolites catalysts comprising anacid 10-membered ring channels are or comprise H-ZSM-5.

DESCRIPTION OF THE FIGURES

FIG. 1 illustrates the process according to the invention.

DETAILED DESCRIPTION OF THE INVENTION

For the purpose of the invention the following definitions are given:

Naphtha is mainly a mixture of straight-chain, branched and cyclicaliphatic hydrocarbons. Naphtha is generally divided into light naphthahaving from 4 to 10 carbon atoms per molecule and heavy naphtha havingfrom 7 to 12 carbons per molecule. Typically, light naphtha containsnaphthenes, such as cyclohexane and methyl-cyclopentane, and linear andbranched paraffins or alkanes, such as hexane and pentane. Light naphthatypically contains 60% to 99% by weight of paraffins and cycloparaffins.

The term “metal free” as used herein means that in the course of thepreparation of the catalyst composition, no metal is willingly added. Itis always possible that some metals (noble or transition) are present aspollution or traces, but it shall be understood that they are present ata very low concentration if not below the detection limit. As apreferred embodiment, a metal-free catalyst composition contains at most50 ppm wt, more preferably at most 5 ppm wt even more preferably at most1 ppm wt of noble metal and less than 0.05 wt % even more preferablyless than 0.01 wt % of transition metals as based on the total weight ofthe catalyst composition.

The term “alkane” or “alkanes” as used herein describes acyclic branchedor unbranched hydrocarbons having the general formula C_(n)H_(2n+2), andtherefore consisting entirely of hydrogen atoms and saturated carbonatoms; see e.g. IUPAC. Compendium of Chemical Terminology, 2nd ed.(1997). The term “alkanes” accordingly describes unbranched alkanes(“normal-paraffins”or “n-paraffins” or “n-alkanes”) and branched alkanes(“iso-paraffins” or “iso-alkanes”) but excludes naphthenes(cycloalkanes).

The term “aromatic hydrocarbons” or “aromatics” relates to cyclicallyconjugated hydrocarbon with a stability (due to derealization) that issignificantly greater than that of a hypothetical localized structure(e.g. Kekule structure). The most common method for determiningaromaticity of a given hydrocarbon is the observation of diatropicity inthe ¹H NMR spectrum.

The terms “naphthenic hydrocarbons” or “naphthenes” or “cycloalkanes” asused herein describes saturated cyclic hydrocarbons.

The term “olefin” as used herein relates to an unsaturated hydrocarboncompound containing at least one carbon-carbon double bond. Preferably,the term “olefins” relates to a mixture comprising two or more selectedfrom ethylene, propylene, butadiene, butylene-1, isobutylene, isoprene,and cyclopentadiene.

The term “LPG” as used herein refers to the well-established acronym forthe term “liquefied petroleum gas”. LPG, as used herein, generallyconsists of a blend of C2-C4 hydrocarbons i.e. a mixture of C2, C3, andC4 hydrocarbons.

One of the petrochemical products which may be produced in the processof the present invention is BTX. The term “BTX” as used herein relatesto a mixture of benzene, toluene, and xylenes.

As used herein, the term “C# hydrocarbons”, wherein “#” is a positiveinteger, is meant to describe all hydrocarbons having # carbon atoms. C#hydrocarbons are sometimes indicated as just “C#”. Moreover, the term“C#+ hydrocarbons” is meant to describe all hydrocarbon molecules having# or more carbon atoms. Accordingly, the term “C5+ hydrocarbons” ismeant to describe a mixture of hydrocarbons having 5 or more carbonatoms. The term “C5+ alkanes” accordingly relates to alkanes having 5 ormore carbon atoms.

The term “zeolite catalyst” refers to a molecular sieve aluminosilicatematerial. Reference herein to a zeolite catalyst having acid 10-memberedring channels is to a zeolite or aluminosilicate having 10-membered ringchannels in one direction, optionally intersected with 8, 9 or10-membered ring channels in another direction.

The terms “comprising”, “comprises” and “comprised of” as used hereinare synonymous with “including”, “includes” or “containing”, “contains”,and are inclusive or open-ended and do not exclude additional,non-recited members, elements or method steps. The terms “comprising”,“comprises” and “comprised of” also include the term “consisting of”.

The recitation of numerical ranges by endpoints includes all integernumbers and, where appropriate, fractions subsumed within that range(e.g. 1 to 5 can include 1, 2, 3, 4 when referring to, for example, anumber of elements, and can also include 1.5, 2, 2.75 and 3.80, whenreferring to, for example, measurements). The recitation of endpointsalso includes the recited endpoint values themselves (e.g. from 1.0 to5.0 includes both 1.0 and 5.0). Any numerical range recited herein isintended to include all sub-ranges subsumed therein.

The particular features, structures, characteristics or embodiments maybe combined in any suitable manner, as would be apparent to a personskilled in the art from this disclosure, in one or more embodiments.

The process of the invention can be operated at high naphtha feedstockconversion, reducing the need for a further cracking process.

Reference is made to FIG. 1. The invention the invention provides aprocess of catalytic conversion by hydrocracking of a paraffinic andnaphthenic hydrocarbons from a naphtha feedstock 1 to propylene, whereinthe process comprises the following steps:

-   -   a) providing a naphtha feedstock 1 containing one or more        paraffins comprising 4 to 10 carbon atoms with preferably no        olefins i.e. an olefin content of less than 1 wt % preferably        less than 0.1 wt % even more preferably less than 0.01 wt %        relating to the total weight of said naphtha feedstock; and    -   b) submitting said naphtha feedstock 1 to a hydrocracking step 3        by contacting said naphtha feedstock 1 with a catalyst        composition in the presence of hydrogen in a reaction zone under        hydrocracking conditions to produce an effluent 5;    -   c) submitting the effluent 5 to a separation step 7 to recover        propane 15; and d) submitting said propane 15 to a step of        dehydrogenation 19 into propylene 23 in a propane        dehydrogenation reactor;    -   e) collecting hydrogen 21 produced in the step of        dehydrogenation into propylene 23, and recycling said hydrogen        21 back to the hydrocracking reaction zone in order to perform        the hydrocracking step 3 and wherein the catalyst composition of        the hydrocracking step 3 comprises one or more zeolite catalysts        comprising an acid 10-membered ring channels.

The effluent 5 resulting from the hydrocracking step 3 can be submittedto a separation step 7, wherein:

-   -   The C4 paraffins fraction 11 can be recycled back to the        hydrocracking reactor to be submitted to a further hydrocracking        step 3.    -   The C5+ hydrocarbons fraction 17 can be valorized as gasoline or        aromatics. Indeed, the C5+ fraction shows a significant        improvement in Research Octane Number (RON) as compared to the        naphtha feedstock 1.    -   The hydrogen 9 can be recycled back to the hydrocracking reactor        to be used in the hydrocracking step 3.    -   The C1 and C2 fraction 13 (methane and ethane fraction) can be        use as off-gas.

The process further comprises a step of collecting hydrogen 21 producedin the step of dehydrogenation into propylene 23, and a step ofrecycling said hydrogen 21 back to the hydrocracking reaction zone inorder to perform the hydrocracking step 3.

With preference, the process comprises a step of recovering theunconverted propane 15 after the step of dehydrogenation 19 intopropylene 23 and recycling it to the propane dehydrogenation reactor tobe submitted to a further step of dehydrogenation 19 into propylene.

The naphtha feedstock 1 used in the invention comprises paraffinic andnaphthenic hydrocarbons, preferably the naphtha feedstock 1 comprisesone or more paraffins comprising 4 to 10 carbon atoms.

The naphtha feedstock 1 may comprise compounds other than paraffins.Preferably, the naphtha feedstock comprises at least 10 wt % ofparaffins comprising 4 to 10 carbon atoms as based on the total weightof the naphtha feedstock, more preferably at least 50 wt %, morepreferably at least 60 wt % of paraffins comprising 4 to 10 carbonatoms.

Preferably, the naphtha feedstock 1 comprises in the range of from 10 wt% to 100 wt % of paraffins comprising 4 to 10 carbon atoms as based onthe total weight of the naphtha feedstock, more preferably of from 50 wt% to 99.5 wt %, more preferably of from 60 wt % to 95 wt % of paraffinscomprising 4 to 10 carbon atoms.

The naphtha feedstock 1 may comprise straight run naphtha or naphthafractions derived from natural gas, natural gas liquids or associatedgas. The naphtha feedstock 1 may comprise naphtha fractions derived frompyrolysis gas. The naphtha feedstock 1 may also comprise naphthas ornaphtha fractions obtained from a Fischer-Tropsch process forsynthesising hydrocarbons from hydrogen and carbon monoxide. Forexample, the naphtha feedstock 1 is or comprises desalted crude oil.

The naphtha feedstock 1 may also comprise higher paraffins, i.e.paraffins comprising more than 10 carbon atoms. Cracking such higherparaffins typically requires the use of temperatures and pressures whichare at the higher end of the preferred temperature and pressure ranges.

Preferably, the naphtha feedstock 1 comprises at least 10% ofnaphthenes. More preferably, the naphtha feedstock 1 comprises in therange of from 10 to 40 wt %, more preferably of from 50 to 90 wt % ofnaphthenes and paraffins C6+, based on the total weight of the naphthafeedstock.

The naphtha feedstock 1 may comprise olefins. However, as the olefinsare hydrogenated during the hydrocracking process, the presence ofolefins results in an undesired increased hydrogen consumption.Preferably, the naphtha feedstock 1 comprises in the range of from 0 to20 wt % of olefins, based on the total weight of the naphtha feedstock,more preferably of from 0 to 10 wt % of olefins. Optionally, the naphthafeedstock 1 is subjected to a hydrogenation treatment prior to beingsupplied to a process according to the present invention.

In an embodiment, the naphtha feedstock 1 comprises at least 2.0 wt % ofaromatics of five or more carbon atoms as based on the total weight ofthe naphtha feedstock, preferably at least 2.2 wt %, more preferably atleast 2.5 wt %.

In an embodiment, the naphtha feedstock 1 comprises at most 10.0 wt % ofaromatics of five or more carbon atoms as based on the total weight ofthe naphtha feedstock, preferably at most 9.0 wt %.

In a preferred embodiment, the hydrocracking step 3 is performed in asingle reactor. Indeed, the invention provides a one-stage hydrocrackingprocess.

The one or more zeolites catalysts having acid 10-membered ring channelsthat can be used for the invention can be selected from:

-   -   one-dimensional zeolites catalysts having 10-membered ring        channels in one direction, which are not intersected by others        channels from another direction;    -   three-dimensional zeolites catalysts having intersecting        channels in at least two directions, whereby the channels in one        direction are 10-membered ring channels, intersected by 8, 9 or        10-membered ring channels in another direction.

Examples of zeolites catalysts having acid 10-membered ring channelssuitable for the process of the invention can be of, but not limited to,the MFI-type, the MEL-type, the MSE-type, the CON-type, the ZSM-8-type,the FER-type, the MTT-type, the TON-type, the EUO-type, the MFS-type,the AEL-type and the ZSM-48-type zeolites catalysts. Preferably, thecatalyst is or comprises a zeolite of the MFI-type. The zeolite arepreferably under their acidic form or H-form MFI-type have athree-dimensional structure, preferably the zeolites catalysts of theMFI-type are selected from ZSM-5, silicalite-1. The preferred MFI-typezeolite is ZSM-5. MEL-type have a three-dimensional structure,preferably the zeolite of the MEL-type is selected from ZSM-11,silicalite-2, and SSZ-46. The preferred zeolite of the MSE-type isMCM-68. The zeolite of the CON-type is selected from CIT-1 and SSZ-33.The zeolite of the FER-type is selected from Ferrierite, FU-9 andZSM-35. The preferred zeolite of the MTT-type is ZSM-23. The zeolite ofthe TON-type is selected from ZSM-22, Theta-1 and NU-10. The zeolite ofthe EUO-type is selected from ZSM-50 and EU-1. The preferred zeolite ofthe MFS-type is ZSM-57. The preferred zeolite of the AEL-type isSAPO-11. ZSM-48 refers to the family of microporous materials consistingof silicon, aluminium, oxygen and optionally boron. All the zeolite arepreferably under their acidic form or H-form

Preferably, the catalyst composition comprises one or more zeolitescatalysts having an acid 10-membered ring channels selected from thelist comprising ZSM-5, silicalite-1, ZSM-11, silicalite-2, SSZ-46,MCM-68, CIT-1, SSZ-33, ZSM-8, Ferrierite, FU-9, ZSM-35, ZSM-23, ZSM-22,Theta-1, NU-10, ZSM-50, EU-1, ZSM-57, SAPO-11 and ZSM-48. Morepreferably the catalyst is or comprises ZSM-5 zeolite. All the zeoliteare preferably under their acidic form or H-form

In a preferred embodiment, the catalyst composition comprises 3D zeolitewithout cages (cavities) and containing at least one 10-member rings.

Preferably, the catalyst composition comprises at least 60 wt % of oneor more zeolite catalysts having an acid 10-membered ring channels, morepreferably at least 70 wt %, even more preferably at least 80 wt % andmost preferably at least 90 wt %.

Several mesoporisation approaches may be used to create zeolitescrystals that contain both mesopores and micropores, includingdestructive approaches such as demetallation (desilication anddealumination) and recrystallization; and constructive approaches suchas using hard templates, supramolecular templates, and surfacesilanization.

To provide sufficient acidity for the hydrocracking reaction, it ispreferred that the zeolites catalysts are at least partly in theirhydrogen form, e.g. H-ZSM-5 or H-ZSM-11. Preferably at least 50 wt % ofthe total amount of the zeolites catalysts used are in their hydrogenform, preferably at least 80 wt %, more preferably at least 90 wt %, andeven more preferably 100 wt % of the zeolites catalysts are in theirhydrogen form.

When the zeolites catalysts are prepared in the presence of an organiccation, they may be activated by heating them in an inert or oxidativeatmosphere to remove the organic cation. For example, they may beactivated at a temperature over 500° C. for at least 1 hour.

In a preferred embodiment, the one or more zeolite catalysts have aSi/Al molar ratio ranging from 10 to 100, preferably from 20 to 80, morepreferably from 30 to 60.

In an embodiment, the zeolite is shaped with a binder. The binder is aninorganic material.

Preferably, the binder is selected from silica, alpha-alumina,gamma-alumina, clays, alumina phosphates, mullite, zirconia, titania,yttria, silicon nitride, silicon carbide, iron, bronze and stainlesssteel, glass, and carbon, preferably the binder is alumina.

The zeolite shaped with the binder forms a catalyst composition, and thecatalyst composition of the present invention preferably comprises atleast 10 wt % of a binder as based on the total weight of the catalystcomposition, most preferably at least 20 wt % of a binder and preferablycomprises up to 40 wt % of a binder.

Optionally, the zeolite may be modified by phosphorous, steaming,leaching, mesoporization, dealumination or modification withalkali-earth or rare-earth metals.

In a further aspect, the catalyst composition used in the process of thepresent invention is prepared by the method comprising the steps of:

-   -   steaming acid zeolite catalyst at a temperature between 500°        C.−750° C. for a period from 0.1 to 24 h under steam pressure        from 0.1 to 10 bars;    -   optionally, contacting the steamed zeolite with the one or more        source of phosphorus;    -   optionally, introducing to the phosphate sample at least 0.1 wt        % of Mg, Ca, Sr, Ba, Ce, La, Fe, Ga;    -   drying and steaming of the one or more modified acid zeolite        catalysts at a temperature between 500° C.-750° C. for a period        from 0.1 to 24 h under steam pressure from 0.1 to 10 bars.

When at least 0.1 wt % of Mg, Ca, Sr, Ba, Ce, La, Fe, Ga is introducedto the phosphate sample, the said metals may be further presented oncatalyst in form of oxides, silicates, aluminates or phosphates.

Accordingly, the one or more acid zeolite catalysts are contacted with asolution in which one or more basic compounds are dissolved, andwherein, with preference, one or more redox elements are dissolved aswell. Preferably, the solution is an aqueous solution. The preferredsource of phosphorous is the phosphoric acid. The preferred solublesalts of the basic compounds and of the redox elements are nitratesalts. The preferred soluble salts of the basic compounds are selectedfrom the list consisting of Mg(NO₃)₂, Ca(NO₃)₂, Sr(NO₃)₂, La(NO₃)₃,Ga(NO₃)₃, Fe(NO₃)₃.

In an embodiment, the phosphorous modified acid zeolite catalyst ispreferably obtained by the process described in WO2009/016156, which isincorporated herein by reference. The process comprises the followingsteps in this order:

-   -   selecting a zeolite with low Si/Al molar ratio (advantageously        lower than 30) among H⁺ or NH₄ ⁺-form of MFI, MEL, FER, MOR,        clinoptilolite, said zeolite having been made preferably without        direct addition of organic template;    -   steaming at a temperature ranging from 400 to 870° C. for 0.01        to 200 h;    -   leaching with an aqueous acid solution containing the source of        P at conditions effective to remove a substantial part of Al        from the zeolite and to introduce at least 0.3 wt % of P;    -   separation of the solid from the liquid;    -   an optional washing step or an optional drying step or an        optional drying step followed by a washing step;    -   a calcination step.

The basic compounds and the optional redox element(s) and phosphorus(P), may be deposited by contacting the one or more acid zeolitecatalysts with a single solution in which the soluble salts of the basiccompounds, soluble salts of the redox elements and phosphoric acid aredissolved.

Alternatively, the basic compounds and the optional redox element(s) andphosphorus (P) may be deposited by subsequently contacting the one ormore acid zeolite catalysts with the different elements and/orphosphorus, whereby the composition is dried to evaporate the solventbefore contacting the composition with the following element. Afterdepositing all the required elements, the resulting composition(catalyst precursor) is dried.

In one embodiment of the present invention, the catalyst precursor isair-dried, preferably for about 8 hours at a temperature ranging from60° C. to 80° C. while stirring.

After drying, the zeolite-comprising composition, on which the basiccompounds and the optional redox element(s) and the phosphorus (P) aredeposited, is calcined in an oxygen-comprising atmosphere, preferably ina moisture-free atmospheric air. Preferably, the catalyst precursor iscalcined at a temperature ranging from 450° C. to 550° C. to remove theresidual amount of nitrates and carbons.

Most preferably, the catalyst precursor is calcined at about 500° C. forabout 4 hours. When a binder is present, it is preferred that the one ormore acid zeolite catalysts are mixed with the binder prior tocontacting the one or more acid zeolite catalysts with one or moresolutions comprising soluble salts of basic compounds and the optionalsoluble salts of redox elements and phosphoric acid.

In the process of the invention, the naphtha feedstock is contacted withthe catalyst composition at elevated temperatures and elevated pressuresin hydrocracking conditions to perform a hydrocracking step.

In an embodiment, the hydrocracking conditions comprise the naphthafeedstock being contacted with the catalyst composition at a temperatureranging from 200° C. to 600° C., preferably ranging from 250° C. to 550°C., more preferably ranging from 300° C. to 450° C. Preferably, thehydrocracking conditions comprise the naphtha feedstock being contactedwith the catalyst composition at a pressure ranging from 1 to 10 MPa,preferably in the range of 2 to 6 MPa, more preferably from 3 to 5 MPa.

In an embodiment, the hydrocracking conditions comprise the naphthafeedstock being contacted with the catalyst composition at a weighthourly space velocity of the naphtha feedstock (VVHSV) of at least 0.1h⁻¹, preferably is ranging from 0.1 h⁻¹ to 10.0 h⁻¹, more preferablyfrom 0.5 h⁻¹ to 8.0 h⁻¹, even more preferably from 1.0 h⁻¹ to 6.0 h⁻¹,and most preferably from 1.5 h⁻¹ to 5.0 h⁻¹.

Hydrogen may be provided at any suitable ratio to the paraffinscontained in the naphtha feedstock. Preferably, the hydrogen is providedin a molar ratio hydrogen to the paraffins in the naphtha feedstockranging from 1000:1 to 1:1, more preferably from 100:1 to 1:1, even morepreferably from 20:1 to 1:1; wherein the number of moles of theparaffins in the naphtha feedstock is based on the average molecularweight of the naphtha feedstock. The process according to the inventionis to achieve a set conversion of the hydrocarbon naphtha feedstock.Preferably, the ratio of hydrogen to the paraffins in the naphthafeedstock is chosen such that the process conditions enable to achievethe desired conversion.

PDH is a catalytic dehydrogenation process that converts paraffins (inthis case propane) to their corresponding light olefins (propylene)dating back from the 1930s. In the late 1980s, the chromia-aluminacatalyst was specifically applied in the dehydrogenation of propane topropylene. The conversion process is favored under high temperature andlow partial pressure of propane. The reaction is run optimally at 500°C.-700° C. to minimize thermal cracking, while the reaction pressure istypically atmospheric. The product mixture goes through a deethaniser toremove light hydrocarbons and traces of hydrogen. The last separationstep involves a propane-propylene splitter to achieve polymer-gradepropylene. It features high yields of 90 wt % propylene, low generationof by-products and relatively low investment and operating costscompared to steam cracking. Lummus CATOFIN and UOP OLEFLEX are the mostcommonly licensed and proven technology. Less common, at the globallevel, are the ThyssenKrupp Udhe STAR process and Linde/BASF's PDHtechnology. Recently, Dow Chemicals introduced a new fluidized bedpropane dehydrogenation technology to the market.

The catalytic dehydrogenation of propane is an endothermic reaction,which produces propane and hydrogen. Hydrogen can be reused for naphthacracking. The extent of conversion is limited by the thermodynamics ofthe reaction with higher temperature favouring higher conversion. Theselectivity decreases as the conversion increases.

In U.S. Pat. No. 6,392,113 B1 assigned to ABB Lummus, the performanceand economics of a catalytic dehydrogenation process are improved byusing two pre-reactors before two main dehydrogenation reactors (one inoperation and the other in regeneration). In this process, propane ispreheated and then passed to the pre-reactor, wherein the catalyst bedis not heated, and propane is partially dehydrogenated with conversionat about 10-25% with the effluent leaves the reactor at 100° C. Thepre-reactor can be operated for hours before it is subjected toregeneration. The partially dehydrogenated reactor effluent is thenreheated by passing through a fired unit, where the heat is supplied bya portion of the effluent air from the regeneration of the maindehydrogenation reactor. The remainder of the effluent air is used toregenerate the catalyst in the pre-reactor. The heated effluent from thepre-reactor is then dehydrogenated in the main dehydrogenation reactor.The operation of the pre-reactor is comparatively steady than that ofthe main reactor, which permits an extended cycle of about 24 hoursbefore any catalyst regeneration is required. In WO95/23123, the overallperformance of the dehydrogenation process can be improved with respectto the dehydrogenation cycle and heating cycle by the change fromco-current to counter-current flow through the catalyst bed. In theprocess, the regeneration gas is introduced at the opposite of thereactor from the feed hydrocarbon. This provides the highest temperatureat the end of the reactor, thus creating the most favourable conditionsfor the dehydrogenation reaction with respect to equilibrium. Thereaction is carried at 590° C. and 0.05 MPa in the presence of chromiumcatalyst supported on alumina with a conversion of propane at 47% andselectivity to propylene at 87%.

In the UOP Oleflex process, the dehydrogenation reaction is carried outat about 600° C. and 0.1 MPa in four sequential moving-bed reactors withinterheaters to reheat the reactor effluent to the desired reactiontemperature before passing it to the next catalyst bed (U.S. Pat. No.5,321,192). Hydrogen and other inert compounds may be added to the feedstream to the dehydrogenation reactor. The catalyst comprises 0.7-0.75wt % Pt, 0.5 wt % Sn, and 3.5-4.4 wt % alkali metal supported onγ-alumina (EP 0448858 B1, U.S. Pat. No. 5,457,256). It passes throughannular bed via gravity flow (U.S. Pat. Nos. 5,227,567, 5,177,293). Thereactor effluent is compressed and dried before it passes to a cryogenicoperation system (a cold box) where hydrogen-methane gas is removed fromC2+ hydrocarbon compounds. The gas is subjected to adsorption with anadsorbent, such as alumina. Silica gel, active carbon, or molecularsieves are used to remove methane from the hydrogen gas (U.S. Pat. No.5,457,256). In U.S. Pat. No. 6,293,999 assigned to UOP, a pressure swingadsorption (PSA) process is used for the separation of a hydrocarbonfeed gas comprising propylene and propane into a fraction, whichcomprises predominantly propylene and a fraction comprising propane. Theprocess uses a small pore aluminophosphate molecular sieve, ALPO-14, toselectively adsorb propylene while essentially excluding propane at anadsorption temperature of 70-120° C. and at a propylene partial pressureof 0.05-0.2 MPa. The desorption conditions are conducted under atemperature of 70-120° C. and a propylene partial pressure of 0.001 and0.05 MPa. Overall, the process comprises the adsorption step, the seriesof connected co-purge steps, and the counter-current depressurizationand repressurization step. In another UOP patent (U.S. Pat. No.6,218,589), a selective hydrogenation is employed to treat a mixture ofthe reactor effluent from a dehydrogenation and a recycle stream from adownstream propane-propylene splitter, and to convert the majority ofhighly unsaturated impurities, such as methylacetylene and propadiene,to propane. The effluent from the selective hydrogenation reactor issent to a deethanizer before being purified within a propylene-propanesplitter.

In a preferred embodiment the step 19 of dehydrogenation of propane 15to propylene 23 (i.e. the PDH step) is performed with a catalystselected from:

-   -   0.7-0.75 wt % Pt, 0.5 wt % Sn, and 3.5-4.4 wt % alkali metal        supported on γ-alumina;    -   Cr₂O₃/Al₂O₃;    -   Pt—Ga/Al₂O₃.

In a preferred embodiment, the PDH step 19 is performed at a temperatureranging from 500 to 800° C. and a partial pressure of propylene belowone atmosphere.

In a preferred embodiment, the PDH step is performed as described inUS2010/0236985. In an embodiment, the catalyst used in thedehydrogenation step of propane to propylene comprises:

-   -   i. from 0.25 to 5.0 wt %, preferably 0.3 to 3.0 wt % of the        first component, preferably gallium, or a compound thereof;    -   ii. from 0.0005 to 0.05 wt %, preferably 0.0007 to 0.04 wt % of        the second component, preferably platinum, or a compound        thereof;    -   iii. from 0.0 to 2.0 wt %, preferably 0.1 to 1.0 wt % of an        alkali metal or alkaline earth metal, preferably potassium; and    -   iv. a support comprising alumina in the gamma crystalline form.

Methods

Gas chromatography was performed on Columns: DB1 (40 m, 0.1 mm, 0.4 μm)and Al₂O₃ (50 m, 0.32 mm, 5 μm) using Agilent operated by ChemStationsoftware.

The elemental composition (i.e. the metal content) of catalystcomposition can preferably be determined by ICP-OES according forinstance to the method UOP Method 961-12. The platinum content canpreferably be determined according to the method ASTM D4642.

EXAMPLES

The following examples illustrate the invention.

Example 1

The process was conducted in a fixed bed reactor loaded with a ZSM-5(Si/Al-40, CBV8014 from Zeolyst INT) containing catalyst extruded withan Al₂O₃ binder (80 wt % zeolite, 20 wt % Al₂O₃) in form of cylinders1.6 mm. After the extrusion, the catalyst composition was dried at roomtemperature for 24 hours followed by calcination at 550° C. for 6 h. Thedemonstration of the invention was performed in both micro pilots.

A stainless-steel reactor tube having an internal diameter of 10 mm isused. 10 ml of the catalyst composition, as pellets of 35-45 mesh, isloaded in the tubular reactor. The void spaces, before and after thecatalyst composition, are filled with SiC granulates of 2 mm. Thetemperature profile is monitored with the aid of a thermocouple placedinside the reactor at the top of the catalyst bed. Prior to thereaction, the catalyst composition was pretreated with an hydrogen flowat 400° C. for 6 h (heating rate 60° C./h) followed by cooling down tothe reaction temperature. The performance test is performed down-flow at4 MPa of total pressure, molar H₂/Naphtha of 13.2; WHSV (naphtha) of 2.5h⁻¹, temperature of 400° C. Analysis of the products is performed byusing an on-line gas chromatography. The results are provided in tables1 to 3.

Feedstock Characteristics

TABLE 1 Naphtha feedstock, POINA analysis nPar iPar Napht Arom Total36.38 38.06 23.28 2.29 C1 0 0 0 0 C2 0 0 0 0 C3 0 0 0 0 C4 8.1 0.58 0 0C5 15.51 14.38 3.9 0 C6 11.43 17.75 14.97 1.98 C7 1.34 5.35 4.41 0.3 C80 0 0 0 C9 0 0 0 0 C10 0 0 0 0 C11 0 0 0 0 >200 C. 0 0 0 0 >200 C. 0 0 00 >200 C. 0 0 0 0 P + A

TABLE 2 The distillation cut (according to ASTM D 86) is given in below:DIST_86: Distillation ASTM D86 T° C. at IBP 28.9° C. DIST_86: T° C. at5% vol 38.1° C. DIST_86: T° C. at 10% vol 40.3° C. DIST_86: T° C. at 20%vol 44.2° C. DIST_86: T° C. at 30% vol 48.4° C. DIST_86: T° C. at 40%vol 52.6° C. DIST_86: T° C. at 50% vol 57.1° C. DIST_86: T° C. at 60%vol 61.6° C. DIST_86: T° C. at 70% vol 66.6° C. DIST_86: T° C. at 80%vol 72.1° C. DIST_86: T° C. at 90% vol 79.4° C. DIST_86: T° C. at 95%vol 89.9° C. DIST_86: T° C. at FBP 91.3° C. DIST_86: % Recovered at76.2% vol 70° C. (*) DIST_86: Recovered % vol 97.0% vol DIST_86: Residue% vol  0.5% vol DIST_86: Loss % vol  2.5% vol DIST_86: (*) % Lossincluded N Yes/No

The density was determined according to ISO 12185 at 15° C. and wasfound to be 0.6702 g/ml.

TABLE 3 FEED EFFLUENT Effluent with C4 (in wt %) (in wt %) recycling (inwt %) Paraffins 74.4 92.8 91.4 Cy-Paraffins 23.3 0.3 0.1 Olefins 0.000.28 0.4 Aromatics 2.3 6.6 8.1 Breakdown Methane 0 1.7 2.3 Ethane 0 4.76.3 Propane 0 38.5 51.6 i-Butane 0.58 13.0 0.2 n-Butane 8.1 12.1 0.2 C5+91.3 30.0 37.1

The process allows valorizing about 50% naphtha as propane and about 37wt % as gasoline, which shows a very high carbon efficiency.

Taking into account an overall propylene yield in propanedehydrogenation of about 82%, the process allows converting of about 41wt % of naphtha to propylene as opposed to 17 wt % in steam cracking.This offers about 2.4 times higher propylene yield from the samefeedstock.

TABLE 4 RON of the C5+ fractions FEED Effluent RON 74 (71-77) 92 (89-93)Aromatic content, wt % 2.5 21.8 Olefins content, wt % <0.1 <1.0

The C5+ produced in the reaction demonstrates very valuable propertiesfor the use of gasoline.

The catalyst composition was two times regenerated in the air at 550° C.and demonstrated full recovering of the activity.

The process shows a low H₂ consumption. Indeed, the naphtha feedstockcontained 15.7 wt % of hydrogen as based on the total weight of thefeed, whereas the weight content of hydrogen in the effluent was 16.4 wt%.

Example 2

The experiment was repeated with compositions comprising zeolitescatalysts comprising 12-membered ring channels. They showed fastdeactivation and significantly lower yield of propane, as shown by theresults given in table 5.

Conditions: pressure: 4 MPa

-   -   Molar H₂/Naphtha: 13.2    -   WHSV (naphtha): 2.5 h−1    -   Temperature: 450° C.

TABLE 5 FEED Composition Catalyst of the feed MOR ZSM-5 BETA ZSM-12SI/AI 10 40 40 40 Structure 1D 3D 3D 1D Members ring 8, 12 10 12 12Composition of the effluent Paraffins 84 78 92.5 79 83 Cy-paraffins13.23 1.15 0.05 6.76 0.30 Olefins 0.00 8.61 0.92 2.69 0.93 aromatics 2.911.8 6.5 11.2 16.0 Conversion to C5− Total — 59.8 83.9 15.6 60.0 Propane0.0 27.4 43.9 5.8 32.0 Ethane 0.0 7.4 12.7 0.6 3.4 Iso-butane 0.1 11.811.6 5.3 12.4 N-butane 2.3 7.6 9.0 2.8 9.4 Conversion to C5+ Total 10040.2 16.1 84.4 40.0 % aromatics 2.9 26.9 40.6 13.3 39.1 in C5+ BTX (inarom) 100 81.3 89.0 75.9 76.5

Many zeolites catalysts deactivate in the presence of aromatics in thefeed. ZSM-5 shows stable performances.

The results demonstrate a higher selectivity for propane than for ethanefor all the zeolite catalysts tested. However, best results wereachieved with ZSM-5.

1.-15. (canceled)
 16. A process of catalytic conversion by hydrocrackingof paraffinic and naphthenic hydrocarbons from a naphtha feedstock topropylene, the process being characterized in that it comprises thefollowing steps: a) providing a naphtha feedstock containing one or moreparaffins comprising 4 to 10 carbon atoms; and b) submitting saidnaphtha feedstock to a hydrocracking step by contacting said naphthafeedstock with a catalyst composition in the presence of hydrogen in areaction zone under hydrocracking conditions to produce an effluent; c)submitting the effluent to a separation step to recover propane; and d)submitting said propane to a step of dehydrogenation into propylene in apropane dehydrogenation reactor; e) collecting hydrogen produced in thestep of dehydrogenation into propylene, and recycling said hydrogen backto the hydrocracking reaction zone in order to perform the hydrocrackingstep; and in that the catalyst composition of the hydrocracking stepcomprises one or more zeolite catalysts comprising one or more acid10-membered ring channels.
 17. The process according to claim 16,characterized in that the naphtha feedstock comprises at least 10 wt %of naphthenes as based on the total weight of the naphtha feedstock. 18.The process according to claim 16, characterized in that the one or morezeolite catalysts have a Si/Al molar ratio ranging from 10 to
 100. 19.The process according to claim 16, characterized in that thehydrocracking conditions of the hydrocracking step comprise: a. thenaphtha feedstock being contacted with the catalyst composition at atemperature ranging from 200 to 600° C., and/or b. the naphtha feedstockbeing contacted with the catalyst composition at a pressure ranging from1 to 10 MPa.
 20. The process according to claim 16, characterized inthat the catalyst composition contains no added noble metal and no addedtransition metals.
 21. The process according to claim 16, characterizedin that the catalyst composition contains at most 50 ppm wt of noblemetal and less than 0.05 wt % of transition metals as based on the totalweight of the catalyst composition.
 22. The process according to claim16, characterized in that the hydrocracking conditions of thehydrocracking step comprise the naphtha feedstock being contacted withthe catalyst composition at a WHSV (feed) of at least 0.1 h⁻¹
 23. Theprocess according to claim 16, characterized in that, in thehydrocracking step, hydrogen is provided to the naphtha feedstock at amolar ratio H₂/Naphtha ranging from 1000:1 to 1:1.
 24. The processaccording to claim 16, characterized in that the catalyst compositioncomprises at least 60 wt % of one or more zeolite catalysts comprisingone or more acid 10-membered ring channels.
 25. The process according toclaim 16, characterized in that the catalyst composition comprises oneor more zeolite catalysts selected from the list comprising ZSM-5,silicalite-1, ZSM-11, silicalite-2, SSZ-46, MCM-68, CIT-1, SSZ-33,ZSM-8, Ferrierite, FU-9, ZSM-35, ZSM-23, ZSM-22, Theta-1, NU-10, ZSM-50,EU-1, ZSM-57, SAPO-11 and ZSM-48.
 26. The process according to claim 16,characterized in that the one or more zeolite catalysts comprising oneor more acid 10-membered ring channels comprise or are zeolitescatalysts of the MFI-type.
 27. The process according to claim 16,characterized in that the catalyst composition comprises a binderselected from silica, alpha-alumina, gamma-alumina, clays, aluminaphosphates, mullite, zirconia, titania, yttria, silicon nitride, siliconcarbide, iron, bronze and stainless steel, glass, and carbon.
 28. Theprocess according to claim 16, characterized in that the step ofdehydrogenation into propylene is performed at a temperature rangingfrom 500 to 800° C. and a partial pressure of propylene below oneatmosphere.
 29. The process according to claim 16, characterized in thatthe separation step comprises recovering a C4 paraffins fraction fromthe effluent of step b) and a step of recycling back the C4 paraffinsfraction to the hydrocracking reaction zone, and/or the separation stepcomprises recovering a C5+ hydrocarbon fraction from the effluent ofstep b) and a step of valorization of said C5+ hydrocarbon fraction asgasoline.
 30. The process according to claim 16, characterized in thatit comprises a step of recovering the unconverted propane after the stepof dehydrogenation into propylene and recycling it to the propanedehydrogenation reactor to be submitted to a further step ofdehydrogenation into propylene.
 31. The use of one or more zeolitecatalysts comprising the one or more acid 10-membered ring channels in aprocess according to claim 16 of catalytic conversion by hydrocrackingof paraffinic and naphthenic hydrocarbons from a naphtha feedstock topropane, and further dehydrogenation of said propane to propylene,characterized in that the zeolite catalysts contains at most 50 ppm wtof noble metal and less than 0.05 wt % of transition metals as based onthe total weight of the catalyst composition.